Process for separating aromatic hydrocarbons



Nov. 20, 1962 F. J. ZUIDERWEG ET AL 3,

PROCESS FOR SEPARATING AROMATIC HYDROCARBONS Filed Oct. 23, 1959INVENTORSI FREDERIK J. ZUIDERWEG GERRIT H. REMAN ALBERT SCHAAFSMA BYJ MyTHEIR ATTORNEY United States Patent 3,ll65,l69 PROCES FOR SEPARATINGARQMATIC HYDRQCARBQNS Frederik I. Zuiderweg and Gerrit H. Roman, both ofAmsterdam, and Albert Schaafsma, The Hague, Netherlands, assignors toShell Oil Company, a corporation of Delaware Filed Qct. 23, 1959, Ser.No. 848,276 Claims priority, application Netherlands Mar. 24, 1959 9Claims. (Cl. 2%8-321) This invention relates to a process for theextraction and recovery of aromatic hydrocarbons from a liquidbydrocarbon mixture.

Several processes for the extraction and recovery of aromatichydrocarbons from liquid hydrocarbon mixtures have already beenproposed. Various high boiling, aromatic selective solvents have beensuggested, among these are diethylene glycol, dipropylene glycol, andsulfolane. Because of the light-heavy selectivity of these solvents, thehydrocarbons in the mixture are extracted in the following order: lightaromatics, heavy aromatics, light paraifins, and heavy paraffins. It hasdeveloped in the use of these high boiling solvents that when theprocess is operated to obtain a high yield of aromatics, contaminationwith light paraflinic material is usually experienced. A procedure forincreasing the purity is described in British Patent No. 739,200 whereina hydrocarbon mix-' ture containing aromatic hydrocarbons is introducedinto a multi-stage extractor at an intermediate point and a glycolic,water-containing selective solvent for aromatic hydrocarbons isintroduced into one end of the extractor. At the same end a raffinatepoor in aromatics and in solvent is withdrawn, while at the other end anaromaticrich extract phase is withdrawn and introduced into acolumnwhich is at a lower pressure than the extractor and wherein only part ofthe most volatile aromatics, together with nonaromatics of equivalentvolatility, are removed as the top products and after condensation arereturned as reflux to the extractor at the end from which the aromaticrich extract is Withdrawn. From theremainder of the extract phase thearomatics are separated from the solvent by distillation in a distillingcolumn at atmospheric or subatmospheric pressure, direct steam beinginjected near the bottom of the distilling column. It is stated in theexample of the British patent that by using this process the purity ofthe recovered aromatics may amount to 97.5% and that this value may beincreased to about 9 8% by using a more complex recovery system for theextract phase comprising three columns instead of two, and using twodifferent reflux streams in stead of one single reflux stream. Thisaddition of a further column together with all its auxiliary equipmentto the recovery system which is obviously a very costly measure, clearydemonstrates the extreme difficulties encountered in attempting toincrease the products purity in extraction processes of the presenttype.

In this connection it should be noted that the specifications fornitration grade toluene require a very high aromatics content, viz. ofat least 98.5% by volume (cf. ASTM specification D84150).

It has now been found that by introducing certain changes in the processas described in British specification No. 739,200 and by adhering tocertain specific operating conditions, it is possible to obtainaromatics having a purity of 97.599% or even higher, while using asimple recovery system that involves two columns only.

The invention will be illustrated with reference to the accompanyingschematic drawing, wherein the sole FIG- URE is a process flow diagramof a preferred embodiment of the improved process.

3,055,169 Patented Nov. 20, 1962 The combination of the followingconditions has been found to give the aforesaid improved results:

(1) The solvent employed is a high boiling, aromatic selective materialhaving an atmospheric (normal) boiling temperature within the range of225 C. to 295 C., and should contain not more than 2% by weight ofwater. Suitable solvents include diethylene glycol, dipropylene glycol,sulfolane, and mixtures thereof.

(2) The temperature in the extraction zone should be between 130 and 155C.

(3) The extract-phase from the extraction system should enter thestripping zone with a temperature between 125 and 155 C.

(4-) The pressure in the stripping zone should be lower than thepressure in the extraction system, but not so low as to causeappreciable flash vaporization (i.e. vaporization by pressure reductionwithout heating) and should be at least 1.5 atmospheres.

(5) The top temperature in this stripping zone should be between 125 and155 C. and the bottom temperature should be at least 165 C.

(6) The condensed top vapors from the stripping zone should be freedfrom substantially all water present as a second liquid phase beforebeing returned to the extrac tion system.

(7) The pressure in the distilling zone should be below 0.5 atmospheres.'1

(8) The bottom temperature in the distilling zo'ne should be at least145 C., but should be at least C: below the bottom temperature in thestripping zone. I i

(9) Direct steam should be injected into said distillingzone in order toenable the hydrocarbons to be distilled off at not too high bottomtemperatures. f

(10) The bottom product from the distilling zone should be cooled by atleast C. to atemperature between 130 and 155 C. before beingreintroduced into the extraction zone. j

Thus, according to the present invention there is provided a process forthe extraction and recovery rot aro-f matic hydrocarbons from a liquidhydrocarbon mixturei containing one or more aromatic hydrocarbons, byintroducing the mixture into a counterflow multi-stage .extractionsystem, introducing, at one end of the extraction: system a highboiling, aromatic selective solvent. having an atmospheric boilingtemperature within the range of 225 to 295 C. and containing dissolvedwater; main-' taining the extraction system under pressure sufflcientto. keep the flowing contents liquid and maintaining the flowingcontents at an elevated temperature; withdrawing a raflinate poor inaromatics and in solvent-from the same end of the system as that atwhich the solvent is introduced; withdrawing an aromatic-rich extractphase from the other end of the system, stripping said extract phase ina stripping zone to liberate a mixture of aromatic and non-aromatichydrocarbons, said stripping being effected at a pressure lower thanthat prevailing in the extraction system; condensing the so liberatedvapors and returning the condensate to the extraction zone; conductingthe remainder of the extract phase to a distilling zone operat ing at asubatmospheric pressure, in which distilling zoneseparation is effectedbetween hydrocarbons and solvent, with direct steam being introducedinto said zone; and returning substantially hydrocarbon-free solventcontaining dissolved water to the extraction system. The temperature inthe extraction system is maintained between and C.; the solvent suppliedto one end of the extraction system contains not more than 2% by weightof Water. tion system is introduced into the stripping zone at a tem-.perature between 125 C. and 155 C. without the occur rence ofappreciable flash vaporization. The stripping The extract-phasewithdrawn from the extrac-.

zone is operated at a pressure of at least 1.5 atmospheres absolute, ata top temperature between 125 C. and 155 C. and at a bottom temperatureof at least 165 C. The vapors from said stripping zone are condensed andrecycled to the extraction system after removing substantially all waterpresent as a second liquid phase. The condensed vapors are introducedinto the extraction system at the end at which the extract phase iswithdrawn and/ or at an intermediate point of the extraction zone. Thebottom product from the stripping zone is passed to the distilling zonewhich is maintained at a pressure of less than 0.5 atmospheres absoluteand at a bottom temperature of at least 145 C., the bottom temperaturebeing at least 10 C. lower than the bottom temperature in the strippingzone. The bottom stream from the distilling zone is cooled by at least15 C. to a temperature between 130 and 155 C. and then reintroduced intothe extraction system.

The present process can be applied to feed stocks having a wide or anarrow boiling range. It is especially suitable for separating aromaticsfrom catalytically reforming gasolines, such as hydroformates andplatformates, or from fractions thereof. The feed stocks should have anASTM final boiling point of not higher than 220 C. Preferably the uppercutting point should be not higher than 160 C. When using a reformatefraction as starting material the upper cutting point shouldadvantageously be approximately the same as the upper cutting point ofthe feed stock for the reforming operation. The lower cutting point ofthe feed stock should preferably be about 100 C.

When carrying out the process according to the invention the volumetricratio of solvent to hydrocarbon feed should be between 3:1 and 8:1,preferably between 4:1 and 621, whereas the ratio of the amount of topproduct returned from the stripping zone to the extraction system to theamount of hydrocarbon feed should be between 0.2:1 and 0.7:1, preferablybetween 03:1 and 05:1.

Various high boiling solvents may be used in the present process. Amongthese are diethylene glycol, dipropylene glycol and sulfolane.Diethylene glycol is preferred. The suitable solvents boil within thetemperature range of 225 C. to 295 C. The solvent used in the extractionsystem may contain a small amount, not exceeding 2% by weight, of water.

The extraction system should be a countercurrent multistage extractionsystem, e.g. a column containing packing material, or sieve plates, arotating disc contactor, a multiplicity of mixer-settler combinations,and the like. The number of theoretical stages should preferably be atleast 5 The feed to the extraction system may be introduced at anintermediate point, but it is generally preferred to introduce it at orat least near that end of the extraction system at which the aromaticrich extract phase is withdrawn, because in that case the bestcompromise between product purity and recovery (yield) is usuallyrealized. Suitable inlet points are at said end of the extraction system(viz. at the first theoretical stage) or at the second theoreticalstage, the latter embodiment being preferred because in that case thepurity is distinctly higher whereas the recovery is only slightlydecreased. Feed introduction at a plurality of points between the middleand the extraction phase end of the system may be useful under specificcircumstances.

The stream of hydrocarbons and solvents that is obtained as top productfrom the stripping zone and (after condensation and after removingsubstantially all water separating as a second liquid phase in thecondensing operation) is returned to the extraction system, should alsobe introduced at one or more points at or near the end of the extractionsystem at which the aromatic-rich extract phase is withdrawn. It ispreferred to introduce this stream at or near the end of the system(i.e. at the first or second theoretical stage of the system) because inthese cases the best compromise between product purity and recovery isrealized. The recycle material may in some instances be advantageouslyreturned to the extraction system in a plurality of streams between themiddle and the extract phase end of the system. If necessary, therecycle stream may be heated before entering the extraction system.

The temperature in the extraction system should be between 130 and 155C. There may be a certain temperature gradient over the system, providedthat the above limits are adhered to.

The extract phase leaving the extraction system is introduced in thestripping zone, at or near its top, with a temperature between and 155C., which implies that there will be no or substantially no intentionalcooling of this stream between extraction system and stripping zone.

The stripping zone is operated at a pressure that is lower than thatprevailing in the extraction system but that is at least 1.5atmospheres. The difference in pressure in the extraction system and inthe stripping zone should be not so large as to cause appreciable flashvaporization. In fact, if there is any vaporous material present in theextract phase when entering the stripping zone, the gravimetric ratio ofthis vaporous material to the fresh feed to the extraction system shouldnot exceed 1:10. Preferably this ratio should be below 5:100, and in themost preferred embodiment no vaporous material will be present at all.In this zone the top temperature should be between 125 and 155 C., andthe bottom temperature should be at least 165, the difference betweentop and bottom temperature being at least 40 C. The top product containssubstantially all water and non-aromatic hydrocarbons present in theextract phase, and, in addition, some solvent and aromatic hydrocarbons.

In order to substantially reduce or prevent deterioration of the solventwhen applying high bottom temperatures in the stripping zone it isgenerally advisable that the solvent should contain a minor proportion,e.g. between 0.0 5 and 1% by weight of phenothiazine or of a substitutedphenothiazine. The use of these compounds for the purpose referred to isdescribed in greater detail in the copending United States applicationSerial No. 862,761, of Philip J. Garner, filed December 30, 1959.

The bottom product leaving the stripping zone is passed to thedistilling zone operating at a pressure below 0.5 atmospheres,preferably 0.2 atmospheres and at a bottom temperature of at least 145C., which temperature should always be at least 10 C. and preferably atleast 20 C. lower than the bottom temperature in the stripping zone.There should be no or substantially no other cooling of the streampassing from the stripping zone to the distilling zone otherwise than byexpansion.

Whereas the stripping zone will be operated without introduction ofdirect steam near or in the bottom part, direct steam should heintroduced into the distilling zone because otherwise the temperaturesrequired to obtain a substantially hydrocarbon free solvent as thebottom product would be so high as to cause decomposition of theglycolic solvent.

As a consequence of this introduction of steam in the distilling zonethe solvent leaving the distilling zone contains a certain amount (lessthan 2% by weight) of dissolved water.

The bottom product from the distilling zone is cooled by at least 15 C.to a temperature between and 155 before being recycled to the extractionsystem.

In the preferred embodiment of the present process the followingoperating conditions are adhered to:

Temperature in extraction system C Temperature of extract phase onentering stripping zone C 128 Top temperature in stripping zone C 128Bottom temperature in stripping zone C 175 Top temperature in distillingzone C 85 Bottom temperature. in distilling zone C 155 Temperature ofbottom stream from distilling zone after cooling C 140 Pressure inextraction system atrn. abs 5 Pressure in stripping zone atm. abs 1.8Pressure in distilling zone atm. abs 0.2

Under these conditions the water content of the solvent leaving thedistilling column will be about 0.6% by Weight.

The present process, in which there is substantially no flashvaporization between extraction system and stripping zone, a ratherlarge temperature drop over the stripping zone, a distinctly higherbottom temperature in the stripping zone than in the distilling zone,and a cooling between distilling zone and extraction zone, isfundamentally different from the approach of U.K. patent specificationNo. 739,200, already mentioned above, in which the temperature drop overthe stripping zone is rather small (as reflected by statement that it isusually not necessary to add heat to the stripping zone by means of areboiler), in which the main heat input in the recovery system isobviously in the bottom of the distilling zone and not in that of thestripping zone, and in which there is obviously no external coolingbetween distilling zone and extraction zone.

As a consequence of this essentially diflerent situation the presentprocess has the advantage that a greater proportion of the non-aromaticsstill present in the extract phase is removed in the overhead of thestripping zone, which results in a higher purity of the final extract ascompared with the previous process.

A liquid hydrocarbon mixture containing both aromatic and non-aromatichydrocarbons, is extracted in a multistage countercurrent extractor 1,which operates under pressure and at a temperature between 130 and 155C. For this purpose the feed is introduced into the extractor throughone or more of the feed inlet lines 2 to 5, whereas the selectivesolvent, containing a small amount of dissolved water, is introducedinto the extractor 1 at or near its top through line 6.

The raflinate phase, which contains only relatively small amounts ofwater, solvent and aromatic hydrocarbons, is withdrawn from the top ofthe extractor 1 through line 7 and further processed to removesubstantially all solvent present therein.

The aromatic-rich extract phase is withdrawn from the bottom of theextractor through line 8 and passes through a reducing valve (not shown)to the stripper column 10, the cooled extract phase being introduced ator near the top of the stripper. The pressure in column 10 is at least1.5 atmospheres, but lower than the pressure in extrac tor 1.

The stripper 10 is provided with a reboiler 11 and is operated with afairly high temperature drop over the column. In the stripper,separation is effected into a top product containing some solvent, partof the aromatic hydrocarbons and most of the water and non-aromatichydrocarbons present in the extract phase leaving the extractor 1, and abottom product that contains only a small amount of water and containssolvent and aromatic hydrocarbons and at most a minor amount ofnon-aromatic hydrocarbons. The top vapors are passed through line 12 tothe condenser 13 and the resulting liquid passed to the settler 14,wherein separation in two layers, viz. a waterrich layer and a layerconsisting mainly of solvent and hydrocarbons, takes place. The waterlayer is removed via line 15, whereas the solvent-hydrocarbon layer ispassed through line 16 and heater 17, to one or more of the lines 18 to20, through which it enters the extractor in its lower part.

The bottom product from the stripper is passed through line 21, andreducing valve (not shown) to distilling column 22. In column 22, whichoperates at subatmospheric pressure, separation is effected into a topproduct 6 containing aromatic hydrocarbons and substantially free fromsolvent, and into a bottom product, that is substantially free fromaromatic hydrocarbons.

The top vapors are withdrawn through line 23, provided with condenser 24and the condensate is partly withdrawn via line 2 5, as thearomatic-rich product, and partly passed through line 26 equipped withthe settler 26A, wherein separation in two layers, viz. a lowerwaterrich layer and an upper layer consisting mainly of hydrocarbons,takes place. The upper layer is recycled as reflux to column 22. Thebottom product from this column is Withdrawn through line 27 andreturned to the extractor through line 6 after cooling in cooler 28. Ifdesired, part of the bottom product may be withdrawn through line 29 forpurification or rejection, and fresh or purified solvent may beintroduced through line 30. Distilling column 22 is operated without areboiler, hot steam being introduced into the bottom of this column vialine 31 and heater 32 to strip dissolved hydrocarbons from thedescending liquid.

It should be realized that the drawing and its above description areschematic and that many auxiliary features such as valves and pumps,have not been represented or discussed.

We claim as our invention:

1. A process for the extraction and recovery of aromatic hydrocarbonsfrom a liquid hydrocarbon mixture containing aromatic and non-aromatichydrocarbons comprising (l) introducing the mixture into a counterflowmulti-stage extraction system wherein the temperature is maintainedbetween and 155 C.; (2) introducing, at one end of the extractionsystem, a high boiling, aromatic selective solvent having a normalboiling tempera ture Within the range of 225 C. to 295 C. and whichsolvent contains dissolved water in the proportion of not more than 2%by weight; (3) passing the solvent in countercurrent flow to thehydrocarbon mixture to obtain a solvent extract phase enriched inaromatics; (4) withdrawing the extract phase from the extraction systemand introducing the said extract phase at a temperature between 125 and155 C. without appreciable flash vaporization into a stripping zone,which is operated at a pressure of at least 1.5 atmospheres absolute andat a top temperature between 125 C. and 155 C. and at a bottomtemperature of at least 165 C. to separate overhead a vapor phaseenriched in non-aromatic hydrocarbons; (5) condensing the vapors fromsaid zone and forming a first liquid phase consisting essentially ofhydrocarbons along with a second liquid phase and recycling them to theextraction system after removing substantialy all water present as thesecond liquid phase, with the condensed vapors being introduced into theextraction system between the intermediate point of the system and theend at which the extract phase is withdrawn; (6) withdrawing the liquidremainder of the extract phase from the stripping zone and separatingthe solvent from the hydrocarbons thereof in a distilling zone at apressure of less than 0.5 atmospheres absolute, said zone having abottom temperature of at least C., which is at least 10 C. lower thanthe bottom temperature in the stripping zone; and (7) withd-rawin g thestripped solvent from the distilling zone, cooling said stripped solventby at least 15 C. to a temperature between 130 and C., and returning thecooled solvent to the extraction system.

2. A process in accordance with claim 1 wherein the aromatic selectivesolvent is selected from the group consisting of diethylene glycol,dipropylene glycol, sulfolane, and mixtures thereof.

3. A process according to claim 1 in which the feed mixture is acatalytically reformed gasoline, or a fraction thereof.

4. A process according to claim 1 in which the feed stock has an uppercutting point not higher than C.

5. A process according to claim 1 in which the feed to the extractionsystem is introduced near that end of the extraction system at which theextract phase is withdrawn.

6. A process according to claim 1 in which the condensed vaporssubstantially free from water are introduced near that end of theextraction system at which the extract phase is withdrawn.

7. A process according to claim 1 in which the extract phase isintroduced in the stripping zone near the top of said zone.

8. A process according to claim 1 in Which the Weight ratio of thevaporous material present in the extract phase on entering the strippingzone to the fresh feed to the extraction system is less than 5: 100.

9. IA process according to claim 8 in Which no vapors are present in theextract phase on entering the stripping zone.

References (Iited in the file of this patent UNITED STATES PATENTSSchumacher Dec. 15, 1942 Durrum Sept. 17, 1946 Weedman Oct. 9, 1956Poffenberger Aug. 20, 1957 Findlay Sept. 10, 1957 Hawkins et al Oct. 8,1957 Broughton Mar. 17, 1959 Schofield Sept. 22, 1959 FOREIGN PATENTSGreat Britain Oct. 26, 1955

1. A PROCESS FOR THE EXTRACTION AND RECOVERY OF AROMATIC HYDROCARBONSFROM A LIQUID HYDROCARBON MIXTURE CONTAINING AROMATIC AND NON-AROMATICHYDROCARBONS COMPRISING (1) INTRODUCING THE MIXTURE INTO A COUNTERFLOWMULTI-STAGE EXTRACTION SYSTEM WHEREIN THE TEMPERATURE IS MAINTAINEDBETWEEN 130 AND 155*C., (2) INTRODUCING, AT ONE END OF THE SYSTEM, AHIGH BOILING AROMATIC SELECTIVE SOLVENT HAVING A NORMAL BOILINGTEMPERATURE WITHIN THE RANGE OF 225*C. TO 295*C. AND WHICH SOLVENTCONTAINS DISSOLVED WATER IN THE PROPORTION OF NOT MORE THAN 2% BYWEIGHT; (3) PASSING THE SOLVENT IN COUMTERCURRENT FLOW TO THEHYDROCARBONS MIXTURE TO OBTAINED A SOLVENT EXTRACT PHASE ENRICHED INAROMATICS; (4) WITHDRAWING THE EXTRACT PHASE FROM THE EXTRACTION SYSTEMAND INTRODUCING THE SAID EXTRACT PHASE AT A TEMPERATURE BETWEEN 125* AND155*C. WITHOUT APPRECIABLE FLASH VAPORIZATION INTO A STRIPPING ZONE,WHICH IS OPERATED AT A PRESSURE OF AT LEAST 1.5 ATMOSPHERES ABSOLUTE ANDAT A TOP TEMPERATURE BETWEEN 125 *C.AND 155*C. AND AT A BOTTOMTEMPERATURE OF AT LEAST 165*C. TO SEPERATE OVERHEAD A VAPOR PHASEENRICHED IN NON-AROMATIC HYDROCARBONS; (5) CONDENSING THE VAPORS FROMSAID ZONE AND FORMING A FIRST